Conversion of hydrocarbon gases into liquid hydrocarbons



Nov. 28, 1939.I c. R. WAGNER CONVERSION 0F HYDROCARBON GASES INTO LIQUID HYDROCARBONS Filled June 17, 1936 2 Sheets-Sheet 1 MSNM MSG.

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OVERHEAD Fgf INVENTQR gary R. Wagner ATTORNEY PatentedNov. 2s, 1939- CONVERSION F HYDROCARBON GASES INTO -LIQUID HYDDOCARBONS Cary B.. Wagner, Chicago, lll., assignor to The Pure Oil Company, Chicago; lll., a corporation of Ohio Application :une 1T, 193s, serial No. sam

llclnlms.

This invention relates to the conversion of gaseous hydrocarbons into liquids .and is more particularly concerned with the conversion of oleilnic hydrocarbon gases into aromatic liquid 5 hydrocarbons.

It is well known that hydrocarbons of the methane series can be converted into aromatic hydrocarbons by Aheating to above 1200* F. and preferably above 1400 F. at atmospheric pressure or at super-atmospheric pressures up to 200 pounds per square inch. It is also known that olefin-containing gases can be converted into hydrocarbons boiling within the gasoline range by heating to temperatures of approximately l5 7001050 F. under highy super-atmospheric pressure. However, the polymers resulting from the latter process are predominantly of the oleflnic or naphthenic type.

'I'he object of this invention ls to convert olenic gases into aromatic hydrocarbons.

Another object of the invention is to provide a process of converting straight chain hydrocarbons into aromatics at much lower temperatures than have been used heretofore.

A further object of the invention is to provide a low pressure method of polymerizing olefin containing gases.

Still another object of the invention is to provide a process capable of producing high yields 30 of toluol.

A still further object of the invention is to produce a novel product which may be fractionated into luseful components or may be used as. a motor fuel.

In accordance with the invention, olefin-containing gases in which the concentration of ole ilns is held within certain limits, preferably 40- 60% by weight, are heated to a temperature of from 11001200 F. for a period of time ranging 40 from 5 to 30 seconds and then quickly chilled to a temperature below reaction temperature. The condensate resulting from the process has an unusually high toluol content. 'I'his condensate is not only extremely valuable for the individual 5 compounds that may be fractionated therefrom, but is also valuable as a motor ,fuel because of its high octane number.

In order to more clearly understand the invention, reference should be had to the follow 50 ing description and the accompanying drawings,

of which,

Fig. l is a diagrammatic elevational view of apparatus suitable for carrying out the invention,

5% Fig. 2 is a graph showing the distillation range of the product made in accordance with the process.

Referring to Fig. 1, wet gas from a vapor phase or other cracking unit is charged through line I by means of compressor 3, which compresses the 5 gases to approximately 150-300 pounds per square inch, preferably about 1'75 pounds per square inch, and then passes cooling coil 5 into accumulator 1. The gases in passing through cooling coil 5 are cooled down to substantially l0 atmospheric temperature and any liquids which condense are discharged from the accumulator 1 through the line 9 into the mix tank I I. The uncondensed gases leave the top, of the accumulator 'I through the line I3. These gases will con- 15- tain from 40-55% of unsaturated hydrocarbons ,(olens) if they result from a vapor phase cracking operation, or a lesser percentage of olens if the gases have been produced in a high pressure cracking operation. The gases may come 20 from any other desirable source, as for example# from a gas cracking operation. From the line I3 the gases pass into the lower portion of the abn sorber I5 wherein they are contacted with a counter-current stream of absorber oil such as mineral seal or light gas oil, which is fed into the upper portion of the absorber through the line I1. The pressure in the absorber may be maintained at approximately 150-200 pounds per square inch and preferably about 165 pounds per square inch. At this pressure the heavier gases such as propane, propylene, butane. and butylene, as well as any heavier fractions which may still be present in the gases, together with some ethane and ethylene, are absorbed by the absorberoil. The remaining unabsorbed gases, which will comprise hydrogen, methane, ethane, ethylene, and possibly some propane and pro- `py1ene, will be discharged from the top of the absorber through the line I9. These gases may be either eliminated from the system through the line 2ll controlled by back pressure regulator 23, or may be charged through the line 25 controlled by valve 21, to be mixed with other gases before charging to the polymerization system. It will be evident that the gases leaving the top of the absorber may be split and part withdrawn from the system and part fed to the polymerization unit. p

The rich oil is withdrawn from the lower part of the absorber through the line 3l controlled by valve 33 and passes into a fixed gas eliminator 35 where the pressure is maintained at approxi- 'mately 100' pounds per square inch. A small portion of the gas, comprising the lighter constituents dissolved in the absorber liquid, escape from the top of the fixed gasl eliminator through the line 31 controlled by valve 39. This gas may be withdrawn from the system and used for any desired purpose, as for example, for fuel, or may be charged to the polymerization system. The rich oil is thenwithdrawn from the xed gas' eliminator through the line 4| by means of pump 43 and passes through heat exchanger 45 and preheater 41 into the upper portion of the still 49. The rich oil in the still 49 is circulated through line 5|, reboiler 53 and line 55 in indirect heat interchange with hot oil which is cir.-4 culated fromV a suitable heater (not shown) through the line 51, reboiler 53, 1ine 59, preheater 41, and back to the heater through line 6|. The hot oil passes in indirect heat interchange with the rich oil in the preheater 41 before being returned to the heater.

In the still 49 therich oil is denuded of its absorbed gas and light gasoline fraction, which leave the top of the still through the line 63. From the line 63 the gases and light gasoline fractions pass through the condenser 65 into the separator 69 from which water and condensed gasoline fractions are withdrawn through lines 1| and 13 respectively. The water present in the condensate is due to the open steam which may be charged into the still 49 through theline 15 to assist in stripping the absorbed gas and gasoline fraction from the absorber oil.' The denudedk oil is returned by means of pump 16' through line 16, heat interchanger 45, cooling coil 18 and line |1 to the'top of the absorber l5. The still 49 and separator 69may be -operated at a pressure of approximately pounds per square inch. From the separator the condensed hydrocarbons pass through the line 13 and line 11 into the mix tank The uncondensed gases, composed chieiiy of ethane, ethylene, propane and propylene, leave the top of the separatorl 69 through the line 19 controlled by valve 8| and may be discharged from the system through line 83 controlled by valve 85, or may be united, in whole or part, with the fresh gases charged to the system through the line I. A small amount of gas may be liberated in the mix tank and this gas leaves the top of the tank through the line 81 and joins the gases leaving the top of separator 69 through the line 19.

Condensate isWithdrawn from the mix tank through the line 89 by means of pump 9| and a portion thereof may be charged into the top of the still 49 through the line 93 as redux. The major portion of the condensate is charged through the line 95, heat exchanger 91, and lines 99, into the upper portion of the fractionatlng tower |0l. The fractionating tower is maintained under a pressure of approximately 180-220 pounds per square inch and preferably about 200 pounds per square inch. Liquid from the bottom of the fractionator is charged through line |03 to reboiler |05 where it is heated and the light fractions, such as butylene, propylene propane, and butane, together with some heavier fractions, are liberated and pass from the top of the reboiler through line |01 back to the intermediate portion of the fractionating tower |0|. The unvaporized condensate, which comprises stabilized casinghead or light gasoline, is withdrawn from the bottom of the reboiler through the line |09, heat exchanger 91, cooling coil ||3, and line H5, to storage.

Uncondensed gases and light vapors leave the top of the fractionating tower |0| through line leave the top of the reux accumulator ||1 and pass through cooling coll ||9 into reux accumulator |2| which is maintained under substantially the same pressure as the fractionating tower |0|. Any condensate which collects vin the reflux accumulator is recycled from the bottom thereof through the line |23 by means of pump |25 into the top of the fractionating tower |0| to act as reflux. The uncondensed gases through the line |21 controlled by back pressure regulator |29, and may be withdrawn from the system through the line |`3| controlled by valve |33, but are preferably charged through the line |35 controlled by valve |31 to admixture with the gases in the line 25. The gases withdrawn from accumulator |2| may have an olefin content of (iO-75%. It will be evident that any mixture of gases desired may be fed to the polymerizationycaustic scrubber through line |44 and the spent reagent removed through line |42. It is desirable to remove the hydrogen sulfide because of its corrosive properties and in addition because it reacts with the oleflnic hydrocarbons present in the gases to form mercaptans and other organic compounds which are deleterious to the resulting product. The desulfurized gas leaves the top of the caustic tank |4| through the line |43 and passes into a gas receiver |45, from the top of which gases are withdrawn' through the line |41 and charged to the heating and reaction coil |49. i

Gases may becharged to the heating and reaction coil |49 at a pressure Iranging from approximately atmospheric to pounds per square inch, and in passing therethrough are heated to a temperature of 11001200 F., and preferably from 11251175 F. The gases passing through the heating and reaction coil |49 are so proportioned that the olefinic content thereof is from 40.60% by volume of the total gas.y 'I'he gas is quickly heated to the desired temperature and the latter portionof the coil serves as a soaking and reaction zone in which the t'emperaturerey mains substantially constant for a period of time ranging from 5-15 seconds. The coil |49 is located in a furnace |5|, heated by means of the burner |53. As shown, the coil is heated primarily by convection heat although radiant heat may be used if desired. Convection heat is preferable for the reason that radiant heat has a tendency to heat the outer surface of the tubes to an excessive degree, with the result that the gases contacting the inner surface of the tubes are heated to a much higher temperature than the gases in the middleportion thereof. Convection heat provides more even heating of the gases in the heating and reaction coil. The coil may be of any desired cross-section, but a coil of 11A inches inside diameter has been found to be satisfactory and this coil should be of suiiicient length to provide a reaction time within the reaction temperature range of approximately 5-30 seconds.

. pump 21,1, ma che top of the .-.tabuizer Isa to act' The gases leave the heating and reaction coil |49 through the line |55 and may either pass through the unheated reaction zone |51 or bypass around the reaction zone |51. The reaction zone |51 may take the form of a coil of enlarged cross-sectional area, as for example, a coil of approximately 31/2-4 linches inside diameter, or it may take the form of a reaction chamber. The reaction zone |51 provides additional time for reaction when necessary. Howeverfif the coil |49 is of sufficient length to give the desired time of reaction, the zone |51 may be by-passed by passing the gases through the line |59 controlled by valve |6I. A valve |63 is provided in the line |55 to enable the gases to by-pass the zone |51. After leaving the reaction coil |49 or the reaction zone |51, in the event the gases are permitted to pass therethrough, the reaction products are immediately chilled to a temperature below reac- 20 tion temperaturev by means of direct contact with a cool hydrocarbon liquid which is injected into the reaction products passing through the line |89. The cooling liquid may be condensate formed in the process, or may be hydrocarbon oil or distillate formed in a separate process. The reaction products are preferably chilled to a temperature of approximately 300600 F. It is important to suddenly chill the products from reaction temperature to below reaction temperature in order to prevent carbon deposition which accompanies slow or gradual cooling of the reaction products. The reaction products may be further cooled to approximately atmospheric temperature by passing through the cooler |61 and then charged into an accumulator |69 which may be maintained at a pressure of approximately atmospheric to 25 pounds per square inch.

1 'I'here is a considerable pressure drop, through the heating and reaction coil, so that the pressure at the outlet of coil |49 may be approximately one half the pressure of the entering gases. Uncondensed products may leave the top of the accumulator |69 through line |1| at a temperature of approximately 10U-120 F. through back pressure regulator |13 and are compressed by the compressor |15 to a pressure of approximately -200 pounds per square inch, preferably about pounds per square inch, and compressed gases pass through a preheater |11, where they may be preheated to a temperature of approximately 250 F., lines |19 and |8|, into the upper portion of the stabilizer |83. 'I'he condensate from the accumulator |69 is withdrawn therefrom through line |95 by means of pump |81 and a portion thereof may be charged through line |65 controlled by valve |9| to act as cooling fluid for the products leaving the reaction coils |49- or reaction zone |51. The remaining `portion may be charged through the line |92 controlled by valve |93 through heat exchanger |95, lines |91 the line 203, heat exchanger '|95, line 205, cooler 201, and sent tostorage. The -uncondensed gases leave the top of the stabilizer |83 through the line 209, pass through cooler 2| and into a reflux accumulator 2|3. Any condensate which collects in the accumulator 2|! is returned from the bottom thereof through line 2|5, by means of the as reflux. The uncondensed gases leave the top of the accumulator 2|3 through the line 2 |9 controlled by back pressure -regulator 22| and may be either withdrawn, in whole or in part, from the system through the line 223 controlled by valve 225, but are preferably recycled to the gas receiver |45 through the line 221 controlled by valve 229. These gases may have an olefin concentration of 3040% and are useful as a diluent for gases richer in olefin concentration than that desired in the gases charged to the polymerization system. l

As an example of the process carried out using the reaction zone |51, gases from the line 25, containing approximately 40% of unsaturates, and gases from the line |35, containing approximately 65% of unsaturates, were mixed so that the resulting gas had an unsaturated content of 50.4% by volume. This gas was charged to the heating and reaction coil |49 at a pressure of 149 pounds per square inch and heated therein to a temperature of 1166" F. 'I'he gases remained inthe reaction section of the coil at this temperature for approximately 10 seconds and were discharged therefrom at a pressure of '73 pounds per square inch and passed into the reaction zone |51. The gases entered the reaction zone at a temperature of approximately 1150 F. and left the zone at 715 F. The time of residence in the zone |51 was approximately 35 seconds. 'I'he products leaving the reaction zone |51 werelimmediately chilled with condensate formed in the system to a temperature of 500 F. and the pres'- sure dropped to 56 pounds per squareinch. A total of .164,960 cu. ft. of gas was processed, of which 85,470 cu. ft. was gas from accum tor |2|. No gas was recycled. A total of 184,100 cu. ft. of processed gas was obtained. A total of 1.'16 barrels of condensate was obtained, of which 5.99 barrels was gasoline of 418 F. end point. A yield of 1.98 gallons of condensate and 1.52 gallons of gasoline per thousand cubic feet of fresh gas were obtained. The gas charged to the polymerization system had a specic gravity of 1.278 (airs-1) and the resulting gas had/a gravity of 1.102, and the unsaturate content was 35.1%. 'I'he A. P. I.

,gravity of the 418 F. end point gasoline was 31.8.

The 418 F. end point material was subjected to Engler distillation and gave the following cuts:

This material was blended fifty-fifty with a motor fuel having an octane number of 41.3 and the octane number of the blend was 11.1. The octane number blending value of the product was therefore 100.9. l

In this particular run practically no coke lformed in the reaction coils |49 but a considerable amount of coke formed in the reaction zone The total condensate produced in this run was fractionated into a number of cuts, as follows:

taken, and Podbielniak analyses of the gases were made, with the following results:

Per- Feed (in- A. P. I. Boiling range of Residual gravity Cut g2g; cut, F. ggg; gas

Lighter than bemEOL,... 8

34.2 Benzol frac 'en 20 1GO-222. 1.0 32.5 Toluol fraction- 24 20G-250. l 0.32 31 Xylol 9 225-325. 31. 43 23.7 Below naphthalene 9 30G-400. 10.28 Naphthalene '(over- 4.5 -17.21 char e 56% solid and 18.24 44% iquid). 14. 36 11.5 Blue color (unknown 3 405-490. 4.62 composition) 2. 92 10.2 Green color (unknown 2 Heavier than blue cut. 0.72

composition). Bottoms (specific grav- 17 g Lgl-U- 3 5 It will be evident from these analyses that both the butylene and the-butane takepart '1n the reto the system, and 90,930 cu. ft. of processed gas were produced, having a specific gravity of 1.120. 8.26 barrels of condensate were produced containing 6.82vbarrels of 418 F. end point material or 3.32 gallons of E. P. condensate per thousand cubic feet of fresh fractionator gas charged to the system. During this run the, maximum temperature obtained in the reaction coils v|49 was l-148 F. and the gases were subjected to reaction temperature for a period of about 9.8 seconds.

'The reaction products leaving the coils M9 were by-passed around the reaction zone 151 and chilled by means of cool condensate to a temperature of 500 F. The pressure at the inlet of the coil 149 was 149 pounds per square inch and at the outlet 73 pounds per square inch. 25 pounds per square inch was maintained on the accumulator |69 and the stabilizer was maintained under a pressure of 185 pounds per square inch. The composite gas (fresh gas plus recycle gas) charged to the polymerization unit had an unsaturate content of 48.2% and the processed gas had an unsaturate content of 37%. The condensate formed in this run had an A. P. I. gravity of 30.5 and produced a 418 F. end point materialhaving an A. P. I. gravity of 35.8. 'Ihis material,

when subjected to Engler distillatiom gave the following results:

Percent of! F.

action to a large extent, and the propane and propylene to a lesser extent.

In a third run, during a period of 24 hours, 123,010 cu. ft. of gas, comprising a mixture of fractionator -gas from thev line 135 and absorber gas from the line 25 were mixed in the proportion of 68.56% fractionator gas and 31.44% of absorber gas and charged to the polymerization system. The maximum temperature in the reaction coils I49 was 1154i F. andthe gases were maintained at reaction temperature for a period of approximately 10.9- seconds. The gases entered the coils M9- at a pressure of 113 pounds per square inch and left at a pressure of `61 pounds per. square inch. The reaction products were by-passed around zone |51 and wrechilled by means of arrester fluid to a temperature of 500 F. No gas was recirculated through the line 221. '1.4 barrels -of condensate were produced contain- ,ing 5.87 barrels of 418 F. end point material, or a yield of 2.00 gallons of end point material per thousand cubic feet of fresh gas. The fresh vgas had a specic gravity of 1.335 and contained 52.4% of unsaturates. The processed gas had an S. G. of 1.081 and containedf 40.3% of unsatf urates. The condensate had an A. P. I. gravity of 28.2- and yielded a 418 F. end point material having an A. P. I. gravity of 34.4. When this material was subjected to Engler distillation, it gave the following results:

Percent of! overhead as shown on Fig. 2. A series of plateaux may be noted occurring at the approximate boiling points 175180,' 229+231 and 27o-280 which coincide respectively with the boiling points of benzol 176, toluol 230v", M-xylol 282, P-xylol 280. The horizontal length of the plateaux is 75 a measurement of thepercentage of these compounds present and in this case indicates the presence of approximately 20% of benzol, 15% of toluol, and a smaller amount of xylol. The large amount of toluol in the product is unusual and makes the process extremely valuable since the demand for toluol is great and for that reason its price is considerably higher than benzol.

In runs two and three practically no coke was formed in the process, demonstrating that any cracking that does take place is not drastic, that is, the cracking does not result in the formation of carbon and hydrogen. Mild cracking undoubtedly does occur since the volume of processed gas exceeds the volume of the gas charged. Moreover, the yields obtained on gases of different gravity and the composition of the fresh and processed gases indicate that the heavier olens, particularly butylene, take a more active part in the reaction than the lighter oleflns. 'I'he run made in which the reaction zone |51 was used demonstrates the fact that gradual cooling of the reaction products tends to produce coke and it therefore should be avoided. When the products were taken directly from the reaction coils |49 and quickly chilled before any substantial drop in temperature occurs, practically no coke is formed. Not only does formation of carbon occur during this period of gradual cooling but heavy tarry materials are formed at the expense of the more desirable benzol, toluol, and xylol fractions.

When the percentage of unsaturates in the fresh charging stock substantially exceeds 60%, the carbon to hydrogen ratio is so high that the dehydrogenation of part of the charge to form aromatic compounds such as benzol and toluol, is accompanied by deposition of carbon. Such a condition is to be avoided in that it renders the process less economic in having to provide for the removal of solid carbon deposits in the reaction zones.

With fresh charging stocks having more than 60% of olefins, the process may be operated at reaction temperatures of from 10751150 F.,pref erably 11001l25 F., without carbon deposition. Under these conditions a smaller proportion of the liquid products will be dehydrogenated to aromatic compounds and a larger proportion will be olens and naphthenes which are of lower commercial value.

The formation of toluol cannot be promoted at temperatures substantially below 1100 F. Below substantially 40% of unsaturates the reaction will proceed at the 11251175 F. temperature range set forth above, but the yield of total liquid products will be relatively low although the percentage of benzol and toluol in the liquid product will be as great as found in the product from gases containing percentages of olefins in the preferred range. An increase in reaction temperature will not increase the liquid yield, but will result in excessive dehydrogenation of part of the charge and will be. accompanied by carbon deposition.

It is essential that the pressure remain relatively low since high pressures tend to suppress an increase in the total volume or a decrease in the average molecular weight of thecomponents of the system undergoing reaction. It is necessary that bodies of low molecularweight and of low carbon to hydrogen ratio such as methane and hydrogen should beallowed to form in order that another portion of the charge may be allowed to form toluol andbenzol of high carbon to hydrogen ratio.

At high pressures the reaction is chiefly polymerization instead of dehydrogenation accompanied by polymerization, thereby making the reaction difficult to control, owing to evolution of heat of polymerization in the early stages, and furthermore, at the high temperatures necessary to produce toluol, decompositionl of the products of prior polymerization result in carbon deposition. The inlet pressures to the reaction coil should not exceed 150-250 poundsr per square inch.

It will be seen that not only has a process been provided for preparing high octane number motor fuels under conditions of temperature and pressure considerably below those heretofore used. but theprocess provides a means for producing large quantities of toluol and other higher boiling aromatic hydrocarbons.

vWhat I claim is:

y l. The process of converting olefin containing gases to liquid hydrocarbons containing a large proportion of aromatics boiling above the boiling point of benzol which comprises separating the gas into a fraction containing in excess of 60% olens and a fraction containing less than 60% olefins, mixing suicient low olefin content gas with the high olefin fraction to bring the olefin content thereof Within the range of 40 to 60%, heating the mixed gas in the substantial absence of non-hydrocarbon gas to a temperature between 1125 and 1175 F. under pressures between atmospheric and 250 pounds per square inch, maintaining the gas within the aforesaid temperature rangefor a period of 5 to 30 seconds, cooling the resulting products and separating the gases from the resulting liquids.

2. Process in accordance with claim 1 which is carried out in the substantial absence of catalysts.

3. Process in accordance with claim 1 in which a portion of the fraction containing less than 60% oleilns is mixed with the high olefin fraction to bring the olefin content thereof within the range of to 60%.

4. Process in accordance with claim 1 in which a portion of the reaction gas is mixed with the high olefin fraction to bring the olefin content thereof within the range of 40 to 60%.

5. Process in accordance with claim 1 in which a portion of the gas fraction having an olefin content below 60% and a portion of the reaction gas are mixed with the high olefin fraction in order to bring the olen content thereof within the range of 40 to 60%.

6. 'I'he process of converting hydrocarbon gases containing approximately 40% to 60% of olefins into liquid hydrocarbons rich in toluol which comprises subjecting said gases in the substantial absence of non-hydrocarbon gas to temperatures Within the limits of approximately 1125 F. to 1175 F. and under super-atmospheric pressure below that at which carbon formation is actively promoted for a period of time ranging from 5 to 30 seconds, cooling the reaction products below reaction temperature and separating the liquid from the gaseous reaction products.

l '7. Process in accordance with claim 6 in which the olens are chiefly butylenes.

8. Process in accordance with claim 6 in which the pressure is not in excess of approximately 250 pounds per square inch.

9. Process in accordance with claim 6 conducted in the absence of a catalyst.

10. The process of converting hydrocarbon gases containing approximately 40% to 60% of 'olenns into liquid hydrocarbons rich 1n toluol which comprises subjecting said gases in the substantial absence of catalysts and of non-hydrocarbons to temperatures of approximately 1125 to 1175 F. under super-atmospheric pressure below that at which carbon formation is actively promoted for a period, of time ranging from 5 to CARY R. WAGNER. 

